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| United States Patent |
6,237,365
|
|
Trebble
|
May 29, 2001
|
Apparatus for and method of separating a hydrocarbon gas into two fractions
and a method of retrofitting an existing cryogenic apparatus
Abstract
Disclosed is an improved cryogenic demethanizer, for separating an inlet
hydrocarbon gas having a mixture of hydrocarbon components into a residual
lighter gas fraction and a heavier liquid fraction. The fractionation
column in the demethanizer has a main body portion and an upper portion
enlarged with respect to the main body portion. A packing which may be in
the form of a plurality of contact trays, or random packing is located in
the upper enlarged portion of the column. The invention may be used to
retrofit existing cryogenic demethanizers, or used in new installations.
| Inventors:
|
Trebble; Mark (De Winton, CA)
|
| Assignee:
|
TransCanada Energy Ltd. (Calgary, CA)
|
| Appl. No.:
|
009242 |
| Filed:
|
January 20, 1998 |
| Current U.S. Class: |
62/621; 62/902 |
| Intern'l Class: |
F25J 003/02 |
| Field of Search: |
62/621,902
|
References Cited
U.S. Patent Documents
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| 3362175 | Jan., 1968 | Burns et al.
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| 3626448 | Dec., 1971 | Crawford et al.
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| 3656311 | Apr., 1972 | Kaiser.
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| 3702541 | Nov., 1972 | Randall et al.
| |
| 3959085 | May., 1976 | De Graff.
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| 4061481 | Dec., 1977 | Campbell et al.
| |
| 4140504 | Feb., 1979 | Campbell et al.
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| 4155729 | May., 1979 | Gray et al.
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| 4157904 | Jun., 1979 | Campbell et al.
| |
| 4171964 | Oct., 1979 | Campbell et al.
| |
| 4272270 | Jun., 1981 | Higgins.
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| 4278457 | Jul., 1981 | Campbell et al.
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| 4444577 | Apr., 1984 | Perez.
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| 4453956 | Jun., 1984 | Fabbri et al.
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| 4453958 | Jun., 1984 | Gulsby et al.
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| 4456461 | Jun., 1984 | Perez.
| |
| 4519824 | May., 1985 | Huebel.
| |
| 4600421 | Jul., 1986 | Kummann.
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| 4617039 | Oct., 1986 | Buck.
| |
| 4657571 | Apr., 1987 | Gazzi.
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| 4687499 | Aug., 1987 | Aghili.
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| 4720293 | Jan., 1988 | Rowles et al.
| |
| 4851020 | Jul., 1989 | Montgomery, IV | 62/621.
|
| 4854955 | Aug., 1989 | Campbell et al.
| |
| 4861360 | Aug., 1989 | Apffel.
| |
| 4869740 | Sep., 1989 | Campbell et al.
| |
| 4889545 | Dec., 1989 | Campbell et al.
| |
| 4895584 | Jan., 1990 | Buck et al. | 62/621.
|
| 5275005 | Jan., 1994 | Campbell et al. | 62/621.
|
| 5372009 | Dec., 1994 | Kaufman et al.
| |
| 5555748 | Sep., 1996 | Campbell et al. | 62/621.
|
| 5566554 | Oct., 1996 | Vivayaraghavan et al.
| |
| 5568737 | Oct., 1996 | Campbell et al.
| |
| 5678424 | Oct., 1997 | Nazar.
| |
| Foreign Patent Documents |
| 1041003 | Oct., 1978 | CA.
| |
| 1048397 | Feb., 1979 | CA.
| |
| 1249769 | Feb., 1989 | CA.
| |
| 1307454 | Sep., 1992 | CA.
| |
| WO 96/40604 | Dec., 1996 | WO.
| |
| WO 98/17609 | Apr., 1998 | WO.
| |
Primary Examiner: Capossela; Ronald
Attorney, Agent or Firm: Bereskin & Parr
Claims
What is claimed is:
1. An apparatus for cryogenically separating an inlet hydrocarbon gas
stream comprising at least methane and ethane into a residue gas stream
comprising a major portion of the methane and a heavier hydrocarbon
fraction comprising principally ethane and other heavier hydrocarbons, the
apparatus comprising:
(1) a main inlet for the hydrocarbon mixture;
(2) a main stream connected to the main inlet and including a first
separator for separating liquid and vapour phases, the first separator
having a liquid phase outlet and a vapour phase outlet;
(3) a liquid phase stream connected to the liquid phase outlet and a vapour
phase stream connected to the vapour phase outlet;
(4) a branch stream connected to the main inlet;
(5) a fractionation column including first, second and third inlets, the
first inlet being connected to the liquid phase stream, the second inlet
being connected to the vapour phase stream and the third inlet being
connected to the branch stream, whereby the flow through the third inlet
has substantially the same composition as the inlet gas, the second inlet
being provided above the first inlet and the third inlet being provided
above the second inlet;
(6) a first outlet means located at the bottom of the fractionation column
for the heavier hydrocarbon fraction and a second outlet means at the top
of the fractionation column for the residue stream;
(7) an outlet conduit connected to the second outlet means;
(8) first means for cooling the branch stream, including a first heat
exchanger means provided between the outlet conduit and the branch stream
for heat exchange therebetween, for cooling the incoming hydrocarbon gas
branch stream and heating the residue gas stream in the outlet conduit;
and
(9) second means for expanding and cooling the main stream.
2. An apparatus as claimed in claim 1, wherein the fractionation column
includes a body portion of relatively small diameter and an upper portion
of relatively large diameter, and a transition section between the body
portion and the upper portion wherein the first inlet is provided towards
the upper end of the body portion, the third inlet is provided in the
upper portion of relatively large diameter, and the second inlet is
provided in one of the transition section and an end of one of the body
portion and the upper portion adjacent the transition section.
3. An apparatus as claimed in claim 2, wherein the fractionation column
includes packing comprising a plurality of trays.
4. An apparatus as claimed in claim 3, wherein the upper portion includes
one to six trays and a relatively small disengagement zone above the trays
for liquid and vapour separation.
5. An apparatus as claimed in claim 4, wherein the diameter of the upper
zone is in the range of ten to seventeen feet.
6. An apparatus as claimed in claim 5, wherein the upper section includes a
disengagement zone having a height in the range of three feet or less.
7. An apparatus as claimed in claim 4, which includes reboiler means
connected to the fractionation column, for reboiling lighter hydrocarbon
fractions from the column, and a further branch stream in which the
reboiler means is located and which is connected between the main inlet
and the inlet of the first separator, for providing heat for the reboiler
means from the inlet hydrocarbon gas stream.
8. An apparatus as claimed in claim 7, wherein the second means for
expanding and cooling includes a second heat exchanger means provided in
the outlet conduit and in the main stream, upstream of the first
separator, for heat exchange between flows in the outlet conduit and the
main stream.
9. An apparatus as claimed in claim 8, wherein the first means for cooling
includes a throttle valve in the branch stream and the second means for
expanding and cooling includes a throttle valve in the liquid phase stream
and an expander in the vapour stream and wherein the outlet conduit
includes a compressor driven by the expander.
10. An apparatus as claimed in claim 9, wherein the first and second heat
exchange means are connected so that the residue gas passes through the
first heat exchange means and at least a portion of the residue gas then
passes through the second heat exchange means.
11. An apparatus as claimed in claim 10, wherein the first heat exchange
means comprises two separate heat exchange elements, and wherein one
portion of the residue gas passes through both heat exchange elements, and
another portion of the residue gas passes through just one heat exchange
element before passing to the second heat exchange means.
12. An apparatus as claimed in claim 9, which includes a third heat
exchange means connected between the main inlet and both of the main
stream and the branch stream, whereby both of the main and branch streams
flow through the third heat exchange means, wherein the third heat
exchange means is located in the outlet conduit for heat exchange between
the residue gas stream and the combined flows of the main stream and the
branch stream.
13. An apparatus as claimed in claim 12, wherein, in the outlet conduit,
the second heat exchange means is downstream from the first heat exchange
means and the third heat exchange means is downstream from the second heat
exchange means.
14. An apparatus as claimed in claim 13, wherein the outlet conduit,
downstream from the first heat exchange means, includes, in the following
order, a fourth heat exchange means, the compressor, an additional
compressor, a third heat exchange means for cooling the compressed residue
gas, and a further connection through the fourth heat exchange means
whereby the compressed gas is cooled and gas passing to the compressor is
heated.
15. An apparatus as claimed in claim 1 or 8, which includes a static mixer,
to which the vapour phase outlet of the first separator and the branch
stream are connected, wherein the second and third inlets to the
fractionation column are combined and are connected to the outlet of the
static mixer, the combined second and third inlets being provided at the
top of the fractionation column, whereby the static mixer causes
contacting between the branch stream and the liquid phase stream prior to
the fractionation column.
16. An apparatus as claimed in claim 15, wherein the static mixer is sized
to provide mass transfer substantially equivalent to one theoretical stage
of contacting in a fractionation column.
17. An apparatus as claimed in claim 15, which includes a second separator
connected to the outlet of the static mixer, which second separator
includes a second liquid phase outlet and a second vapour phase outlet,
wherein the second liquid phase outlet is connected to the combined second
and third inlets and the second vapour phase outlet is connected to the
outlet conduit, whereby a portion of the residue gas stream does not pass
through the fractionation thereby to reduce the flow of residue gas in the
fractionation column.
18. An apparatus as claimed in claim 1, wherein the second means for
expanding and cooling the main stream includes an expander connected to
the vapour phase outlet of the first separator and having an expander
outlet for cooled and expanded gas, and wherein a second separator is
connected to the expander outlet, the second separator including a second
liquid phase outlet and a second vapour phase outlet, with the second
liquid phase outlet being connected to the second inlet of the
fractionation column and the second vapour phase outlet being connected to
the outlet conduit, whereby a portion of the residue gas stream does not
pass through the fractionation column thereby to reduce the flow of
residue gas in the fractionation column.
19. An apparatus as claimed in claim 1, wherein the branch stream comprises
largely a liquid phase.
20. A method of retrofitting an existing cryogenic apparatus for separating
a compressed inlet hydrocarbon gas stream comprising at least methane and
ethane into a residue fraction comprising a major portion of the methane
and a heavier hydrocarbon fraction comprising principally ethane and other
heavier hydrocarbons, said existing apparatus comprising:
(a) a main inlet for the hydrocarbon gas;
(b) means for expanding and cooling the inlet gas into a mixture of liquid
and vapour phases connected to the main inlet;
(c) means for separating said liquid and vapour phases comprising:
(i) a fractionation column having a body portion and an upper portion above
the body portion and generally enlarged with respect to the body portion,
said upper portion being substantially empty and being originally intended
to provide a disengagement zone;
(ii) at least one inlet means for supplying the vapour and liquid phases to
the column provided on at least the body portion and connected to the
means for expanding and cooling the inlet gas;
(iii) a first outlet means located at the bottom of the bottom portion for
withdrawing the heavier hydrocarbon fraction from the column;
(iv) a second outlet means located at the top of the upper portion for
withdrawing the residue stream from the column; and
(v) packing means in said body portion for increasing the amount of contact
between the liquid and vapour phases; the method comprising:
(1) providing additional packing in the disengagement zone in the upper
enlarged portion of the column to provide additional contact between the
liquid and vapour phases;
(2) providing a branch stream between the main inlet and the upper portion;
and
(3) providing a first heat exchange means in the branch stream, for cooling
the branch stream so that at least a portion of the branch stream is in
the liquid phase, whereby hydrocarbons discharging into the upper portion
from the branch stream are at least partially liquid.
21. A method as claimed in claim 20, wherein the additional packing in the
upper portion of the column comprises a plurality of trays spaced
vertically in said upper portion of the column.
22. A method as claimed in claim 21, which comprises providing one to six
trays.
23. A method as claimed in claim 22, which comprises providing three trays
into an upper portion which is approximately 121/2 feet in diameter and
approximately 10 feet high, and wherein the trays are positioned in the
upper portions so as to leave a disengagement zone of less than three
feet.
24. A method as claimed in claim 23, wherein the upper portion includes a
frustro-conical portion, wherein one of the trays is provided extending
into the frustro-conical portion.
25. A method as claimed in claim 24, wherein the method is carried out in
an apparatus including a main heat exchange means in the means for
expanding and cooling the inlet gas, which main heat exchange means is
connected to the second outlet means, whereby the residue stream absorbs
heat from the incoming hydrocarbon gas, the method comprising providing
the first heat exchange means in the branch stream as first and second
separate heat exchange elements and connecting the heat exchange elements
such that the residue stream passes through the first heat exchange
element, and only a portion of the residue stream passes through the
second heat exchange element, with the remainder of the residue stream
passing through the main heat exchange means, the portions of the residue
stream subsequently being combined for discharge from the apparatus.
26. A method as claimed in claim 20, the method additionally comprising:
(1) providing means for expanding and cooling said vapour phase of the
inlet gas;
(2) providing a static mixer, to which the branch stream, downstream from
the first heat exchange means and the vapour phase, downstream from the
means for expanding and cooling, are connected, and providing a connection
between an outlet of the static mixer and the upper portion of the
fractionation column.
27. A method as claimed in claim 26, which additionally includes:
(1) providing second separation means, connected to the outlet of the
static mixer and including a second vapour phase outlet and a second
vapour phase inlet;
(2) providing a connection between the second liquid phase outlet and the
upper portion of the fractionation column; and
(3) providing an outlet conduit for residue gas, and connecting the outlet
conduit to both the second outlet means and the second vapour phase
outlet.
28. A method of separating a hydrocarbon feed gas stream comprising at
least methane and ethane into a residue gas fraction comprising a major
portion of the methane and a heavier hydrocarbon fraction comprising
principally ethane and other heavier hydrocarbons, the method comprising:
(1) passing the hydrocarbon gas stream through a first heat exchange means
to cool the gas stream;
(2) separating the gas stream into first, second and third streams, with
the second stream comprising a major portion of the gas flow;
(3) expanding the gas, after it has been cooled, to lower the temperature
of the first, second and third streams, wherein the third stream, after
cooling and expansion, is substantially in the liquid phase;
(4) providing a fractionation column including packing;
(5) supplying the third stream to the top of the fractionation column, the
second stream to the fractionation column below the third stream and the
first stream to the fractionation column below the second stream;
(6) collecting the residue gas fraction from the top of the fractionation
column and the heavier hydrocarbon fraction from the bottom of the
fractionation column, and passing the residue gas fraction through the
first heat exchange means to transfer heat to the residue gas.
29. A method as claimed in claim 28, which includes providing a
fractionation column comprising a lower portion having a relatively small
diameter, an upper portion having a relatively large diameter, and a
transition section therebetween and a packing within the lower and upper
portions, wherein the packing comprises, for the upper portion, six or
less trays; wherein
step (5) comprises supplying the third stream to the top of the upper
portion, the second stream to the top of one of the transition section and
one of the upper and lower portions adjacent the transition section and
the first stream to the lower portion below the second stream; and wherein
step (6) comprises collecting the residue gas fraction from the top of the
upper portion and the heavier hydrocarbon fraction from the bottom of the
lower portion, wherein the residue gas fraction passes through the first
heat exchange means to reheat the residue gas fraction and cool the
incoming gas.
30. A method as claimed in claim 29, which comprises:
splitting the inlet hydrocarbon gas stream into a main stream and a branch
stream, the branch stream comprising the third stream and the main stream
being subsequently split into the first and second streams.
31. A method as claimed in claim 30, which includes splitting a portion of
the hydrocarbon feed gas stream off into a further branch stream, and
passing the further branch stream through reboiling means and then
recombining the further branch stream with the main stream, the reboiler
means, reboiling lighter hydrocarbon fractions from the fractionation
column.
32. A method as claimed in claim 31, which includes passing the main gas
stream through a second heat exchange means and the main and branch
streams together through a third heat exchange means, to cool the main and
branch streams, and passing the residue gas stream through the second and
third heat exchangers to extract heat from the main and branch streams.
33. A method as claimed in claim 32, which further comprises:
passing the main gas stream, after cooling, to a first separator and
separating the main gas stream into liquid and vapour phases, with the
liquid phase forming the first stream and the vapour phase forming the
second stream.
34. A method as claimed in claim 33, which includes passing the first
stream through a throttle valve to expand the first stream and passing the
second stream through an expander, to expand the second stream.
35. A method as claimed in claim 34, which includes passing the residue
stream, after leaving the third heat exchange means, through compressor
and driving the compressor by the expander.
36. A method as claimed in claim 34, which includes combining the vapour
stream and the branch stream in a static mixer, and then supplying the
combined branch and vapour stream as a single stream to the top of the
upper portion of the fractionation column.
37. A method as claimed in claim 36, which comprises mixing together the
branch stream and the vapour stream sufficiently to effect contacting
approximately equivalent to one theoretical stage of contacting.
38. A method as claimed in claim 36, which includes:
(1) separating the combined vapour stream and branch stream in a separator
into a second vapour phase and a second liquid phase;
(2) supplying the second liquid phase to the upper portion of the
fractionation column, and combining the second vapour stream with gas flow
from the top of the upper portion of the fractionation column, to form the
residue gas fraction.
39. A method as claimed in claim 28, which includes:
(1) splitting the inlet hydrocarbon gas stream into a main stream and a
branch stream, the branch stream comprising the third stream;
(2) passing the main gas stream, after cooling, to a first separator and
separating the main gas stream into a first liquid phase and a first
vapour phase, with the first liquid phase forming the first stream and the
first vapour phase forming the second stream;
(3) passing the second stream through an expander to expand the second
stream;
(4) combining the second stream and the branch stream together in a static
mixer, to form a combined stream;
(5) separating the combined stream in a separator into a second vapour
phase and a second liquid phase;
(6) supplying the second liquid phase to the top of the fractionation
column and combining the second vapour phase with gas from the top of the
fractionation column, to form the residue gas fraction.
40. A method as claimed in claim 28, wherein the inlet gas is supplied at a
rate up to 600-1000 MMSCFD.
Description
FIELD OF THE INVENTION
The present invention relates to an apparatus for and a method of
separating or fractionating a hydrocarbon gas, comprising at least methane
and ethane, into a residue gas comprising mainly methane and a heavier
fraction comprising principally ethane and other heavy hydrocarbons, and
to a method of retrofitting an existing cryogenic plant or apparatus so as
to be capable of carrying out the method. The gas will usually be natural
gas, received as either a gas or a liquid, and including ethane, propanes,
butanes and higher or heavier hydrocarbons. More specifically, the
invention relates to an improved fractionation column to separate ethane
and heavier components, often referred to as "ethane plus" cryogenically
from methane and heavier hydrocarbons. The column may be used in new
facilities, or may be used as a retrofit with existing facilities.
BACKGROUND TO THE INVENTION
Cryogenic techniques and apparatus to separate methane and/or ethane (as
well as other lighter gases such as carbon dioxide and nitrogen) from a
gas containing a mixture of hydrocarbon gases have long been known. Such
an apparatus is commonly called a "demethanizer". While the term
"demethanizer" will be used throughout, it is to be understood that the
apparatus and methods described may also be used in other applications
whereby the demethanizer would be operated as a deethanizer. Typically,
the purpose of a gas processing facility is to receive a gas from a
transmission line, efficiently cool and depressurize the gas, extract the
more valuable heavier components (ethane and heavier hydrocarbons,
referred to as "ethane plus"), reheat and recompress the gas, and feed it
back into the transmission line.
Typically, in such an operation, an inlet gas is introduced into a process
facility at a high pressure. The gas is then allowed to expand and cool in
various stages, and the liquid and gas fractions are then separated into
different streams. After several stages of expansion and cooling, the
various hydrocarbon streams are introduced into a fractionation column, or
demethanizer, at different heights in the column. The methane gas is then
separated from the heavier components of the gas, with the methane
component exiting the fractionation column through the top of the column
as a residue gas, and the heavier components exiting the column at its
lower portion (collected as a liquid).
In the fractionation column, there is typically located a packing, which
may be in the form of a number of contact trays. However, any suitable
packing or construction can be used that promotes contact between the
vapour and liquid flows. Thus, conventional Raschig rings or other
standard packing materials can be used in the column. This packing is
designed to increase mass transfer contact between the falling liquids and
rising gases within the column, which increases the efficiency of
liquid-gas separation in the fractionation column. As well, there is
usually located at the upper portion of the fractionation column an
enlarged empty "disengagement section". Typically, the stream entering the
fractionation column at its upper portion (in the disengagement section)
comprises between about 5% and about 50% liquid phase when it enters the
fractionation column. The disengagement section allows the liquids
entering the fractionation column space to separate or "deentrain" from
the vapours with which it is mixed. The empty disengagement section is
designed to alleviate potential problems of carryover of hydrocarbon
liquids into the demethanizer overhead stream (which is supposed to be
vapour). This results in a loss of hydrocarbon liquid product.
Conventionally, the disengagement section is of a large diameter (about 10
to 18 feet) and is about ten to fifteen feet in height.
It has also been known that, to increase the efficiency of a standard
demethanizer, to increase the amount of heavier hydrocarbons removed,
additional packing or contact trays may be added in a separate column,
which is connected with the fractionation column. In order to obtain a
useful increase in efficiency, however, it has been thought that at least
eight additional trays, and as many as twenty additional contact trays
were required in this separate column. Therefore, the modification of
existing facilities to obtain an increase in efficiency has previously
been very expensive, since a whole new column (and all of the additional
equipment associated with it) was required.
One known process for separating different components in a hydrocarbon gas
stream is disclosed in U.S. Pat. No. 4,278,457 (to Campbell and assigned
to The Ortloff Corporation). The claimed improvement in that patent is to
divide the feed gas stream into two separate streams, one of which is
cooled and then depressured through an expander, while the other of which
is cooled to a greater degree and is then depressured through a simple
expansion valve. The separate first and second streams are then supplied
to the fractionation column at different feed points.
This U.S. patent discloses a number of different examples, which either
operate at a flow rate of 6588 lb moles/hr or at a quarter of this flow
rate, i.e. 1647 lb moles/hr. Interestingly, in three examples, where the
input gas flow is separated into three separate inputs, and where a
substantial portion, as much as 76%, of the input gas flow, passes either
into the middle or bottom inlet of the column, the lower flow rate of 1647
lb moles/hr is given. Moreover, this patent does not discuss, in any way,
details of the column design, in terms of diameter required at different
heights, number of trays or number of trays between different inlets.
Even more particularly, in this Ortloff patent, there is no detailed
direction as to the number of trays that might be required between a
middle inlet and a top inlet to a fractionation column, where the major
portion of the gas is supplied to the middle inlet, and a small portion is
provided, substantially in the liquid phase, to the top inlet.
Supplying a substantial portion of the inlet gas to a middle inlet of the
column gives a number of problems. In particular, this has a significant
impact on the dimensions of the column. By far the largest portion of the
feed gas is methane, which in the column travels upwards as a vapour. When
a major portion of the supply gas is fed to one inlet, the column, above
that inlet, will need to have a sufficient diameter to accommodate the
upward flow of methane and heavier components in the vapour phase. To keep
the methane velocities reasonable, and in known manner to prevent
entrainment of the heavier component liquid droplets to be carried
overhead out of the column with the methane, requires a large diameter for
the column. Clearly, this diameter or cross-section will be related to the
intended flow rate through the apparatus.
This requirement for a relatively large cross-section where there is a
large flow of methane in the vapour phase, has often resulted in a second
column being provided, as mentioned above. Thus, in view of the
conventional teaching that a significant number of trays would need to be
provided in any such enlarged upper section, this would often result in a
top section for the column that was simply too large to be supported on
top of the lower section. For this reason, such a section was often
provided as a separate column.
Other hydrocarbon gas separation techniques and apparatus can be found in
U.S. Pat. No. 3,702,541 (Randall et al.), U.S. Pat. No. 4,519,824 (Huebel)
and U.S. Pat. No. 5,566,554 (Vijayaraghavan et al.). All of these patents
disclose relatively complex techniques. What is also striking about all of
these three proposals is that no portion of the inlet gas stream is taken
off and fed separately to the fractionating column. Rather, it appears
that the inventors, in all of these cases, have assumed that it is
advantageous to achieve some separation of the inlet gas, before feeding
this into the fractionation column. The assumption appears to be that if
there is some initial separation, e.g. liquid/vapour separation, then this
will improve the overall performance of the system. Thus, all of these
proposals provide at least one separator in which vapour and gas phases
are separated.
The Randall et al. patent is of interest, since the inlet gas flow is split
into three separate streams, by way of liquid and vapour separators, to
provide three separate inlets to the fractionation column, these being
provided at various levels in the lower stripping section of the column.
The column here is provided with an upper rectifying section of larger
diameter. Gas taken off from the top of the column is separated into
liquid and vapour fractions, and the liquid fraction is pumped back up to
the top of the column, and fed into the top of the rectifying section.
Again, there is no discussion of the size problems where a substantial
portion of inlet gas is fed in below the top of the fractionation column,
resulting in a substantial flow upwards through the top part of the
column. This problem becomes particularly acute, where a plant is designed
for large flow rates. Thus, the present invention is intended to provide a
plant or apparatus suitable for a flow rate as large as a thousand (1000)
MMSCFD, approximately equivalent to 116500 lb moles/hr, i.e. a flow rate
that is an order of magnitude or more greater than that in some of the
prior art proposals discussed above. In conventional design practices, for
such a large flow rate, the upper part of the fractionation column can
require diameters approaching 20 feet, and it is impractical to support
such large diameter sections on top of much smaller diameter lower
sections.
Another common characteristic of all of these earlier proposals is that
there is no detailed investigation or consideration of the behaviour at
the top of the fractionation column. Conventional teaching is that a
substantial disengagement section needs to be provided above the top
trays, to ensure adequate and complete separation of vapour and liquid
phases, so as to ensure that liquid droplets are not carried over by the
vapour flow leaving the top of the fractionation column. As detailed
below, on an actual plant implementation of this invention, the size of
the disengagement space was reduced from 101/2 to 21/2 feet and no liquid
carry over has been observed during several months of operation.
SUMMARY OF THE INVENTION
In accordance with one aspect of the present invention, there is provided
an apparatus for cryogenically separating an inlet hydrocarbon gas stream
comprising at least methane and ethane into a residue gas stream
comprising a major portion of the methane and a heavier hydrocarbon
fraction comprising principally ethane and other heavier hydrocarbons, the
apparatus comprising:
(1) a main inlet for the hydrocarbon mixture;
(2) a main stream connected to the main inlet and including a separator for
separating liquid and vapour phases, the separator having a liquid phase
outlet and a vapour phase outlet;
(3) a liquid phase stream connected to the liquid phase outlet and a vapour
phase stream connected to the vapour phase outlet;
(4) a branch stream connected to the main inlet;
(5) a fractionation column including first, second and third inlets, the
first inlet being connected to the liquid phase stream, the second inlet
being connected to the vapour phase stream and the third inlet being
connected to the branch stream, whereby the flow through the third inlet
has substantially the same composition as the inlet gas, the second inlet
being provided above the first inlet and the third inlet being provided
above the second inlet;
(6) a first outlet means located at the bottom of the fractionation column
for the heavier hydrocarbon fraction and a second outlet means at the top
of the fractionation column for the residue stream;
(7) an outlet conduit connected to the second outlet means;
(8) first means for cooling the branch stream, including a first heat
exchanger means provided between the outlet conduit and the branch stream
for heat exchange therebetween, for cooling the incoming hydrocarbon gas
branch stream and heating the residue gas stream in the outlet conduit;
and
(9) second means for expanding and cooling the main stream.
Preferably, the fractionation column includes a body portion of relatively
small diameter and an upper portion of relatively large diameter, and a
transition section between the body portion and the upper portion, and the
first inlet is provided towards the upper end of the body portion, the
third inlet is provided in the upper portion of relatively large diameter,
and the second inlet is provided in the transition section.
The fractionation column preferably includes packing comprising a plurality
of trays. An important aspect of the present invention is the discovery
that these trays can provide a tray efficiency much greater than usual. As
such only a small number of trays is required, and there could be from one
to six trays in the upper portion. Also, a disengagement zone much smaller
than that usually provided can be used, as it has also been discovered
that this still gives adequate separation and disengagement.
Another aspect of the present invention is directed to retrofitting an
existing plant and provides a method of retrofitting an existing cryogenic
apparatus for separating a compressed inlet hydrocarbon gas stream
comprising at least methane and ethane into a residue fraction comprising
a major portion of the methane and a heavier hydrocarbon fraction
comprising principally ethane and other heavier hydrocarbons, said
existing apparatus comprising:
(a) a main inlet for the hydrocarbon gas;
(b) means for expanding and cooling the inlet gas into a mixture of liquid
and vapour phases connected to the main inlet;
(c) means for separating said liquid and vapour phases comprising:
(i) a fractionation column having a body portion and an upper portion above
the body portion and generally enlarged with respect to the body portion,
said upper portion being substantially empty and being originally intended
to provide a disengagement zone;
(ii) at least one inlet means for supplying the vapour and liquid phases to
the column provided on at least the body portion and connected to the
means for expanding and cooling the inlet gas;
(iii) a first outlet means located at the bottom of the bottom portion for
withdrawing the heavier hydrocarbon fraction from the column;
(iv) a second outlet means located at the top of the upper portion for
withdrawing the residue stream from the column; and
(v) packing means in said body portion for increasing the amount of contact
between the liquid and vapour phases; the method comprising:
(1) providing additional packing in the disengagement zone in the upper
enlarged portion of the column to provide additional contact between the
liquid and vapour phases;
(2) providing a branch stream between the main inlet and the upper portion;
and
(3) a first heat exchange means in the branch stream, for cooling the
branch stream so that at least a portion of the branch stream is in the
liquid phase, whereby hydrocarbons discharging into the upper portion from
the branch stream are at least partially liquid.
Again the additional packing in the upper portion of the column can
comprise a plurality of trays, e.g. one to six, spaced vertically in said
upper portion of the column. There can be three trays provided in an upper
portion which is approximately 121/2 feet in diameter and approximately 10
feet high, and the trays are positioned in the upper portions so as to
leave a disengagement zone of less than three feet.
The upper portion can include a frustro-conical portion, and in this case
one of the trays is provided extending into the frustro-conical portion.
Yet another aspect of the present invention provides a method of separating
a hydrocarbon feed gas stream comprising at least methane and ethane into
a residue fraction comprising a major portion of the methane and a heavier
hydrocarbon fraction comprising principally ethane and other heavier
hydrocarbons, the method comprising:
(1) passing the hydrocarbon gas stream through a first heat exchange means
to cool the gas stream;
(2) separating the gas stream into first, second and third streams, with
the second stream comprising a major portion of the gas flow;
(3) expanding the gas, after it has been cooled, to lower the temperature
of the first, second and third streams, wherein the third stream, after
cooling and expansion, is substantially in the liquid phase;
(4) providing a fractionation column comprising a lower portion having a
relatively small diameter, an upper portion having a relatively large
diameter, and a transition section therebetween and a packing within the
lower and upper portions, wherein the packing comprises, for the upper
portion, a small number of trays;
(5) supplying the third stream to the top of the upper portion, the second
stream to the top of one of the transition section and one of the upper
and lower portions adjacent the transition section and the first stream to
the lower portion below the second stream;
(6) collecting the residue gas stream from the top of the upper portion and
the heavier hydrocarbon fraction from the bottom of the lower portion,
wherein the residue gas stream passes through the first heat exchange
means to reheat the residue gas stream and cool the incoming gas.
Preferably, the method comprises: splitting the inlet hydrocarbon gas
stream into a main stream and a branch stream, the branch stream
comprising the third stream and the main stream being subsequently split
into the first and second streams. Furthermore, a portion of the
hydrocarbon feed gas stream can be split off into a further branch stream,
and the further branch stream is then passed through reboiling means and
then recombining the further branch stream with the main stream, the
reboiler means, reboiling lighter hydrocarbon fractions from the
fractionation column.
One embodiment of both the method and the apparatus of the present
invention includes a static mixer, to which the vapour phase outlet of the
expander and the branch or third stream are connected, wherein the second
and third inlets to the fractionation column are combined and are
connected to the outlet of the static mixer, the combined second and third
inlets being provided at the top of the fractionation column, whereby the
static mixer causes contacting between the branch stream and the liquid
phase stream prior to the fractionation column.
Advantageously, the static mixer is sized to provide mass transfer
substantially equivalent to one theoretical stage of contacting in a
fractionation column.
In some instances, a separator may be utilized to separate the two phase
stream leaving the expander. In this type of process configuration, the
vapours from the separator are not directed to the demethanizer column but
join directly with the overhead vapours leaving the demethanizer column.
The liquids from the separator are generally sent as the top feed to the
demethanizer column. This type of process results in a smaller column
diameter but has the requirement of an additional separator vessel.
Retrofitting this type of process is somewhat different since it is
generally less effective, with respect to ethane recovery, to install
trays in the overhead disengagement section. The static mixer approach is
then useful since the third stream mentioned above, which is essentially
condensed inlet gas, can be mixed with the expander outlet and sent
through a static mixing device upstream of the separator vessel. This
results in substantial additional cooling and absorption of the expander
outlet stream and produces significantly more liquid feed to the
demethanizer column which results in higher recoveries of ethane and
heavier components.
BRIEF DESCRIPTION OF THE DRAWINGS
These and other advantages of the present invention will be more fully and
completely understood, when the following detailed description of the
preferred embodiment is read in connection with the following drawings, in
which:
FIG. 1 is a schematic drawing of a conventional apparatus for separating
methane from a hydrocarbon gas;
FIG. 2 is a schematic drawing of an apparatus in accordance with one aspect
of the present invention;
FIG. 3 is a drawing of a theoretical model of an apparatus, similar to the
apparatus of FIG. 1;
FIG. 4 is a drawing of a theoretical model of an apparatus, similar to the
apparatus of FIG. 2;
FIG. 5a is a drawing of a theoretical model of an apparatus in accordance
with another aspect of the present invention;
FIG. 5b is a drawing of a variant of the theoretical model of the apparatus
shown in FIG. 5a;
FIG. 5c is a drawing of a theoretical model which includes a static mixer
for a plant with a second separator which separates expander vapours and
bypasses them around the demethanizer column;
FIG. 6 is an elevational view of the top part of the fractionation column
of FIG. 4, showing trays; and
FIG. 7 is a schematic plan view showing the trays of FIG. 6.
DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENT
Referring to FIG. 1, there is shown schematically a typical conventional
cryogenic separation plant, referred to generally by reference numeral 10.
The plant 10, has an expansion and cooling area 15, a fractionation column
demethanizer 20, and a recompression area 25. While a conventional
demethanizer will be described here in general terms, it is to be
understood that this is not to be considered limiting to the present
invention, which may be used with any existing conventional demethanizer,
or as a new installation. Additionally, the process conditions associated
with the demethanizer will also be described only generally, since they
are not limiting, and any person skilled in the art will understand all of
the equipment and conditions, and how they may be modified if desired.
A compressed inlet gas which may comprise methane, ethane, propane and
heavier hydrocarbons, as well as smaller amounts of carbon dioxide,
nitrogen and other gases, enters the plant 10 into the expansion area 15,
through an inlet 30, where it is divided into two streams 35 and 40. The
inlet gas may be at a temperature of about 65.degree. F. and at a pressure
of about 400 to 1200 p.s.i.a.
The stream 35 then enters a heat exchanger 45, where the gas is cooled
through heat exchange, for example to a temperature of about -85.degree.
F. The stream 40, which is split off from stream 35 before entering heat
exchanger 45, is directed through a series of heat exchangers 50, 55 and
60. The heat exchangers 50, 55 and 60 provide reboiling for the
fractionation column 20, which is required to maintain the methane content
of the ethane plus liquid recovered typically below 2.0 mole %. This will
be discussed below. Correspondingly, the heat exchangers 50, 55 and 60
cool the incoming gas to a temperature of approximately -75.degree. F. and
a pressure somewhat lower than the inlet pressure, solely due to pressure
losses in the various pipes, heat exchangers, etc.
The streams 35 and 40 are then recombined downstream from the heat
exchanger 45, into a combined stream 62. The stream 62 is then directed
into a low temperature separator 65, where the liquid and gas phases are
separated. The pressure in the separator 65 is approximately 550-1150
p.s.i.a. at a temperature of -85.degree. F. The vapour phase leaves the
low temperature separator 65 as an overhead stream 70, while the liquid
phase leaves separator 65 through a bottom stream 75. The low temperature
separator 65 is conventional, and is well known to those skilled in the
art.
The vapour stream 70 is then depressured and further cooled in an expander
80, and subsequently directed as a stream 85 into the fractionation column
20 at point 90. The bottom stream 75 is typically reduced in pressure by
passing through a valve (not shown), which causes flash evaporation or
expansion to occur.
The fractionation column 20 has a main body 95, and an upper enlarged
portion 100, and the inlet point 90, for the stream 85, is provided
towards the bottom of the enlarged portion 100. The upper enlarged portion
100 may be described as generally conical or "belled" in shape. The design
of the fractionation column 20 is conventional, including the upper
enlarged portion 100, and is known to those skilled in the art. In the
conventional design, the upper belled portion 100 of the fractionation
column 20 is empty, and is designed to have a larger diameter than the
main body 95. This area is known as a disengagement zone, and is provided
to allow the hydrocarbon mixture entering the fractionation column 20 as
the stream 85 (which is largely a vapour) an area where liquids entrained
in the vapour phase may disengage from the vapour phase. As previously
stated, this vapour phase may contain from about 5% to about 50% liquid
content. The upper enlarged portion 100 is often between about 10 to 15
feet in height, and about 10 to 18 feet in diameter, depending on the
volume of gas being processed.
The liquid portion exiting the low temperature separator 65 through the
stream 75 is normally directed to a lower point 105 in the fractionation
column 20, but above the location where reboiling streams 107, 108 and 109
leave the fractionation column 20 and are connected with heat exchangers
60, 55 and 50, respectively. This is typically near the upper portion of
the main body 95 of the fractionation column 20, although this may vary
from system to system. This reboiling system is required to maintain the
methane content of the ethane plus stream leaving the fractionation column
below 2.0 mole %. As will be appreciated, portions of the liquid phase
falling through the fractionation column 20 are redirected out of the
column 20, through the reboiling streams 107, 108 and 109, and are then
passed through the heat exchangers 60, 55 and 50, respectively. This aids
in the cooling of the inlet gas passing through these heat exchangers. The
stream 40, now a cooled stream 40b, is recombined with stream 35 and are
directed to the low temperature separator 65.
Within the main body 95 of the fractionation column 20, there is normally
located a "packing", which may be in the form of a series of "trays" or
contact plates (typical to those trays seen in FIGS. 6 and 7 as will be
later discussed). The contact plates are designed to increase the contact
between the liquid and the vapour phases in the fractionation column,
which in turn increases the efficiency of the ethane and higher
hydrocarbon recovery. In a typical installation, there would be between 15
and 30 trays at a spacing of 2 feet between adjacent trays. As noted
above, any suitable packing, such as Raschig rings can be used instead of
trays.
In the fractionation column 20, the vapour phase (mainly the methane
content of the inlet gas) leaves the fractionation column 20 as an outlet
stream 115, where it is directed through the heat exchanger 45, to aid in
cooling to the inlet gas stream 35. The stream 115 is then recompressed,
for example in a brake compressor 120, and is then further compressed to
pipeline pressure utilizing a large compressor 125 and the compression
section is generally indicated as 25. The stream is then passed through an
aftercooler 130 to lower the residue gas temperature to a level suitable
for reentry into a gas transmission line; as is known, for some
applications, it may be possible to omit the aftercooler 130.
The more valuable heavier, liquid phase of the inlet gas (for example
ethane and heavier) exits through the lower end of the column 20 as a
stream 135, and is collected and stored appropriately.
All of the features thus described are conventional, and will be readily
known and understood by persons skilled in the art.
Referring to FIG. 2, in the present invention, the plant 10 is again shown
schematically and modified by (1) the addition of a reflux section 150,
which allows the reflux of the vapours exiting the upper belled portion
100 of the fractionation column, and (2) the addition of trays or contact
plates in the upper belled portion 100 of the fractionation column. These
additional features provide increased efficiency of recovery of the
heavier hydrocarbon components (ie. ethane and heavier).
For simplicity and brevity parts which are common between the apparatus of
FIG. 1 and the apparatus of FIG. 2 are given the same reference numerals
and description of these common components is not repeated.
In the present invention, the inlet gas stream 30 is initially split into
three streams 35, 40 and 160 after entering the apparatus. The additional
stream 160 is directed generally to the reflux section 150, and more
specifically, through an additional heat exchanger 165, in which the
stream 160 is cooled to a temperature of about -135.degree. F. The heat
exchanger 165 is conventional, and may be similar to the heat exchanger
45. Again, some form of throttle or expansion valve (not shown) would be
provided, so as to reduce the pressure and cause flash expansion of the
hydrocarbon stream.
After being cooled in the heat exchanger 165, the gas is expanded by an
expansion valve (not shown) and sent directly to the fractionation column
20 as a stream 170, where it enters the fractionation column 20 at 172,
near the top of the upper belled portion 100.
In the upper belled portion 100 of the fractionation column 20, there is
located additional packing, which may be in the form of a series of
additional trays or contact plates 226-228 (FIG. 6). The additional trays
226-228 may be similar to the trays located in the main body portion 95 of
the fractionation column 20. These additional trays 226-228 are installed
so as to leave a small disengagement zone between the upper most tray 228
and the upper limit of the fractionation column 20. It is at this location
where the stream 170 enters the upper belled portion 100. The
disengagement zone in the present invention may be as little as about two
feet in height, in contrast to the prior art devices, which required a
much larger disengagement zone, usually on the order of between about 10
to 15 feet in height. It has been surprisingly found that only a few
additional trays 226-228 are required in the upper portion 100 of the
column 20, to achieve a substantial increase in the recovery of ethane and
heavier components from the column. It is believed that this is due to an
unexpectedly and unusual high tray efficiency, which should be above 60,
and is expected to exceed 80. In practice, the use of as few as three
additional trays 226-228 has been sufficient to significantly increase the
recovery, as will be discussed below.
The residue gas leaving the fractionation column 20 through stream 115 is
then redirected through the heat exchanger 165, to cool the inlet gas
passing through this heat exchanger. The residue gas is then split into
two streams 180 and 185. The majority of the residue gas, of the order of
70 to 90%, is sent via stream 180 to the heat exchanger 45, also to aid in
the cooling of the inlet gas stream entering this heat exchanger. The
remainder of the residue gas exits the heat exchanger 165 as stream 185.
The stream 180, after exiting the heat exchanger 45 is recombined with the
stream 185, to form stream 190.
The combined stream 190 is then recompressed and cooled, as previously
described, prior to exiting the process.
It was previously believed that it was unacceptable to include the
additional contact plates in the upper belled portion 100, since it was
believed that there would be insufficient room for the incoming liquid and
vapour phases to separate from each other. It was believed that this would
result in substantial liquid carryover and loss of hydrocarbon liquid
through the upper stream 115 leaving the fractionation column 20. This
would clearly be undesirable. Additionally, given that the upper enlarged
portion 100 of the column 20 in existing facilities is of limited size
(usually between 10 to 15 feet in height), it was believed that there was
insufficient room to add an acceptable number of trays to obtain an
increase in ethane plus recovery which was commercially acceptable. As
previously mentioned, it was believed that a minimum of eight additional
trays would be required to ensure sufficient mass transfer contact between
the rising vapours and the falling liquids. This was due largely to the
assumption that several theoretical stages of contact were required and
that tray efficiencies would be quite low (e.g. 40-60%).
Therefore, if it was desired to increase the efficiency of the
fractionation column 20, and in particular when it was desired to add a
reflux section, such as section 150, to allow reflux of the residue gas, a
separate contactor column would be constructed on the ground. In contrast,
the present invention provides additional trays in the fractionation
column. Therefore, such an installation would require a separate vessel
between about 24 and 60 feet in height and a diameter of about 16 to 17
feet. Given the size of the separate vessel, it would not be feasible to
place it on top of the fractionation column.
The cost of a second vessel is significant, and therefore, may often not be
commercially feasible. Additionally, the extra equipment associated with
the separate column, for example, a cryogenic pump to deliver the liquids
from the second separator to the top of the demethanizer column, would
significantly increase capital and operating costs.
In accordance with the present invention, it has surprisingly been found
that the use of as few as three of these contact trays in the upper belled
portion 100 of the fractionation column 20 can achieve almost the same
level of efficiency as the use of eight to twenty trays, which requires a
separate vessel, but at a much reduced cost. Surprisingly, a high tray
efficiency of 82% has been obtained in a plant performance test, and this
requires a much smaller number of trays to obtain the desired separation.
Conventional design practices allow a tray efficiency of 45-60%. The fact
that so few trays are required makes the size of the upper belled portion
sufficiently small so that it can be supported by the main body 95 of the
fractionation column 20.
In an unoptimized actual plant test, it was found that the present
invention, through modification of an existing facility, increased ethane
recovery by about 8.4%, from about 77% to about 85.4%, and propane
recovery was increased from about 98% to about 98.9%. These increased
efficiencies are each significant, particularly given the relatively
minimal cost and time required to modify an existing system to achieve the
benefit. In an optimized trial basis it is expected that the ethane
recovery will increase by 11%. In simulated comparisons where a grass
roots or new facility is proposed, ethane recovery was increased by about
11% as compared to a conventional design as depicted in FIG. 1.
The present invention enables the volume of the residue gas stream 115 to
be reduced, because of the increased volume of liquid ethane and other
substances removed from stream 85. In an actual plant test, using the same
residue gas compressor horsepower, this enabled inlet gas rates to be
increased from 489 million standard cubic feet per day to 501 million
standard cubic feet per day.
This aspect of the invention will be better understood by reference to FIG.
3. This shows a theoretical schematic of the apparatus of FIG. 1. This has
been analyzed using conventional software, Hysim, a software program
licensed by Hyprotech Ltd. from Calgary. The overall layout is similar to
FIG. 1, with the exception of the items outlined below. Otherwise, like
components are given the same reference numeral and their description is
not repeated.
In FIG. 3, the stream 40 passes through a heat exchanger 140 and exchanges
heat with ethane plus recovered from the column. It then flows through the
heat exchanger 50.
For the heat exchanger 50, a stream 141 is taken from the bottom of the
column 20, passed through the heat exchanger 50 and then back to the
fractionation column 20. The liquid from the fractionation column is
connected to a stream 146 for liquid, which is delivered to a pump 148.
The vessel 142 is shown for simulation purposes only. In practice the
vessel 142 is part of the bottom of the fractionation column 20. The pump
148 discharges the recovered ethane and other products through line 149,
which passes through the heat exchanger 140.
The heat exchanger 45 of FIGS. 1 and 2 is now configured as heat exchangers
167a, 167b. As shown, the stream 115 splits and then recombines prior to
entering compressor 125 for flow through the heat exchanger 167a, 167b.
The combined stream 115 flows to compressors 120, 125. Further, similar to
the FIG. 2 embodiment, the line 40 and the line 35b combine, before
entering the separation vessel 65.
The separation vessel 65 has the outlet stream 75 for liquid, connected to
the fractionation column 20 as before. The vapour line 70 is provided with
a main branch stream 71 connected through the expander 80 and then stream
73 to the top of the fractionation column 20. A bypass stream 72 is
provided although as detailed below, it will often not be used. This
enables the expander 80 to be bypassed, so that it can be serviced without
shutting the whole processing plant down.
FIG. 4 shows a theoretical model of the plant or apparatus of FIG. 2.
Again, elements or components already identified and described are given
the same reference numeral and description of them is not repeated, for
simplicity and brevity.
As before, the principal difference between FIGS. 3 and 4, as for FIGS. 1
and 2, is the inclusion of a separate line 170 discharging into the top
end of the column 20. Here, the expansion valve 162 for the line 170 is
shown. Additionally, the heat exchanger 165 is shown as two separate heat
exchange elements 165a and 165b. As for FIG. 3, two separate heat exchange
elements 167a and 167b are shown, approximately corresponding to the heat
exchanger 45.
For these two theoretical models, in FIGS. 3 and 4, theoretical performance
results have been obtained, for temperature, pressure, molar flow rate and
vapour fraction. These are set out in the following Tables 1 and 2.
TABLE 1
Temperature Pressure Molar Flow Vapour
Location .degree. F. p.s.i.a lb moles/hr Fraction
INLET STREAM 65.12 766 53650.57 1
30
STREAM 40 65.12 766 18666.8 1
exit from heat 64.87 762 18666.8 1
exchanger 166
exit from heat 53.5 758 18666.8 1
exchanger 140
exit from heat -73.04 746 18666.8 .93
exchanger 60
STREAM 35a 64.81 761 34983.77 1
stream 35b -88.18 754 34983.77 .74
stream 62 -84.5 746 53650.57 .82
stream 70 -84.5 746 44160.64 1
stream 75 -84.5 746 9489.93 0
(upstream valve
76)
stream 75 -132.68 316.22 9489.93 .39
(downstream
valve 76)
stream 71 -85.47 736 44160.64 1
(upstream
expander 80)
stream 73 -144.59 290 44160.64 .85
(upstream valve
74)
stream 73 -140.57 316.22 44160.64 .85
(downstream
valve 74)
STREAM 115 -145.05 286 49907.72 1
heat exchanger 53.82 278 8889.79 1
167a exit
(line 115)
heat exchanger 35.44 276 41017.93 1
167b exit
(line 115)
compressor 120 - 38.7 276 49907.72 1
entry (stream 190)
compressor 120 - 70.46 322.1 49907.72 1
exit (stream 190)
compressor 125 - 69.47 306 49907.72 1
entry
compressor 125 - 242.16 823 49907.72 1
exit
STREAM 141 22.02 290 4876.1 0
(upstream heat
exchanger 50)
stream 141 30.66 290 4876.1 0.23
(downstream heat
exchanger 50)
stream 144 30.66 290 1133.71 1
stream 146 30.66 290 3742.39 0
stream 149 - entry 32.34 384.17 3742.39 0
to exchanger 140
stream 149 - exit 53.6 379.17 3742.39 .01
to exchanger 140
TABLE 2
Temperature Pressure Molar Flow Vapour
Location .degree. F. p.s.i.a lb moles/hr Fraction
INLET STREAM 65.06 771 55044 1
30
STREAM 40 65.06 771 16741.16 1
exit from heat 64.93 769 16741.16
exchanger 166
exit from heat 52.38 765 16741.16 1
exchanger 140
exit from heat -80.06 753 16741.16 .87
exchanger 60
STREAM 35a 64.87 768 31763.96 1
stream 35b -77.08 760.2 31763.96 .9
stream 62 -78.54 753 48505.13 0.89
stream 70 -78.54 753 43066.88 1
stream 75 -78.54 753 5438.24 0
(upstream valve
76)
stream 75 -128.6 301.23 5438.24 .39
(downstream
valve 76)
stream 71 -79.5 743 43066.88 1
(upstream
expander 80)
stream 73 -138.16 301.23 43066.88 .87
(upstream valve
74)
stream 73 -138.16 301.23 43066.88 .87
(downstream
valve 74)
STREAM 160 65.06 771 6538.87 1
stream 170 -147.74 757 6538.87 1
(upstream valve
162)
stream 170 -156.36 300.7 6538.87 0.06
(downstream
valve 162)
STREAM 115 150.14 294 50877.09 1
stream 185a -111.76 289 4941.21 1
stream 185b 60.79 283.5 4941.21 1
stream 180 -111.76 289 45929.88 1
heat exchanger 46.32 281 8995.2 1
167a exit
heat exchanger 40.72 279 36934.68 1
167b exit
compressor 120 - 43.61 279 50871.09 1
entry
compressor 120 - 73.77 324.4 50871.09 1
exit
compressor 125 - 73.08 313 50871.09 1
entry
compressor 125 - 243.32 823 50871.09 1
exit
stream 141 17.03 298 5313.04 0
(upstream heat
exchanger 50)
stream 141 27.54 298 5313.04 .21
(downstream heat
exchanger 50)
stream 144 27.54 298 1139.81 1
stream 146 27.54 298 4173.23 0
stream 149 31.56 524.63 4173.23 0
stream 149 52 519.63 4173.23 0
(downstream heat
exchanger 140)
A review of the values given in these two tables will show many close
similarities, which might be expected, in view of the very similar flow
rates. The significant differences are in the conditions and flow rates
for the streams entering the fractionation column. Thus, in Table 1, the
stream 75 enters at a temperature of -132.68.degree. F. and a flow rate of
9489.93 lb moles/hr and stream 73 at a temperature of -144.59.degree. F.
and a flow rate of 44,160.64 lb moles/hr.
When modified to provide the third stream, as indicated in Table 2, the
flow rate for stream 73, is decreased slightly to 43,066.88 at a
temperature of -138.16.degree. F. The bottom stream 75 is reduced to a
greater extent to 5438.24 lb moles/hr and a slightly higher temperature of
-128.6.degree. F. However, there is now the additional top stream 170
which is introduced at a temperature of -156.36.degree. F. and a flow rate
of 6538.87. More significantly, while the vapour fractions for stream 73,
75, differ little between the two examples, stream 170 is introduced
almost entirely in the liquid phase, with a vapour fraction of just 0.06.
The effect of this is to provide, at the top three trays, an upward flow
of methane, originating principally from stream 73, which meets a downward
flow of liquid from stream 170, which is introduced in the liquid phase at
a significantly lower temperature. The effect of this is to create a
downward flow from the top three trays, which would ensure that a
significant portion of ethane and other heavier hydrocarbons are absorbed
from stream 73 and carried down through the column, and are not carried
upwards with the methane gas.
Reference will now be made to FIG. 5a, which shows schematically an
apparatus or plant in accordance with a second aspect of the present
invention. This has been designed as a complete new plant or facility,
rather than as a modification to an existing facility. However, again for
simplicity and brevity, parts, common with earlier Figures, are given the
same reference numerals, and description of these common components is not
repeated.
Here, the heat exchangers 50, 55, 60 are represented by the two sides of
the heat exchangers, as heat exchange elements 50a, 50b and 55a, 55b, 60a,
60b for the respective reboiling streams 109, 108, 107. To separately
identify the inward end and outward end of the stream 40, this is
designated as 40a for the stream flowing to the heat exchangers, and
stream 40b leaving the heat exchangers.
The heat exchanger configuration in this embodiment is somewhat different.
Here, broadly corresponding to the heat exchanger 45 of FIG. 1, there are
two heat exchange elements 200, 202 through which the stream 35 flows. The
stream for the returned residue gas is again indicated as 115, and after
passage through a heat exchanger 206, the stream is indicated as 185. Due
to the different configuration here, the designation 185 is used for the
residue gas stream through to the compressor 120 and compressor 125. The
additional heat exchanger 206 is provided, for heat exchange with the
stream 160 passing to the top of the fractionation column 20.
The inlet stream 30 passes through a valve 208 to a stream 210, connected
to the heat exchange elements 200, 202. Between the heat exchange elements
200, 202, the additional stream 160 is branched off, and passes through
the heat exchanger 206 to the top of the column 20.
To enable the various streams to be identified at different points,
primarily for identifying stream conditions as detailed below, various
suffixes are used. Thus, the stream 160 is identified as 160a and 160b
before and after the heat exchanger 206. This stream 160 also includes a
throttle expansion valve 162, to cause flash expansion. The stream 185 is
variously labelled as 185a, b, c, etc to identify the portions indicated
on FIG. 4. For the streams from the separator 65, the liquid stream is
identified as 75b after expansion through a valve 76, and the vapour
stream is identified as 70 and 73 before and after expansion in the
expander 80.
The plant of FIG. 5a is intended to handle 1.0 billion cubic feet of gas
per day (BCFD), which enters through stream 30 at 613 p.s.i.a and
68.degree. F. from a main gas transmission line. As before, the purpose of
the apparatus is to receive gas from a transmission line, to efficiently
cool and depressure the gas, to extract the valuable heavier components
("ethane plus", i.e. ethane and heavier hydrocarbons), and then to
recompress the gas back into the transmission line. The inlet gas
primarily comprises methane although it contains other species including
carbon dioxide (0.2-1.5%), nitrogen (0.5-1.5%), ethane (3-8%), propane
(0.5-2%), and heavier hydrocarbons (0.5-5% ). The objective of such
apparatus is to provide an extraction facility to remove ethane, propane,
and heavier hydrocarbons in substantive quantities for subsequent resale.
The FIG. 5a apparatus is intended to extract over 60% of the ethane and
over 99% of the heavier propane plus components.
A portion of the inlet gas is split off and sent through the side heat
exchangers 50, 55, 60 on the distillation column 220 to provide reboiling.
The remainder of the inlet gas is sent through exchanger element 200 where
it is chilled to -36.degree. F. using residue gas in line 185. Upon
exiting exchanger 200, the inlet gas is once again split and a portion of
the gas is sent through heat exchanger 206, where 84% of the stream is
liquified. This stream then passes through the valve 162 and enters the
top of the distillation or fractionation column 220. The remainder or bulk
of the inlet gas passes through a second inlet gas exchanger 202 and is
chilled to -72.degree. F. This gas is then mixed with stream 40b and the
combined flow is sent to separator 65. Vapours leaving separator 65 are
expanded adiabatically in expander 80 from 595 p.s.i.a to 304 p.s.i.a
causing the gas temperature to drop from -75.degree. F. to -126.degree. F.
The expander 80 generates power from this expansion and the power is
utilized in driving the brake compressor 120. Liquids from the separator
65 are sent to the distillation column through the valve 76.
The column, here indicated as 220, comprises two sections or portions: a
smaller bottom section or portion 222 primarily dedicated to providing
reboiling, and a larger top section or portion 224 which accomplishes a
majority of the ethane recovery out of the inlet gas. The top section 224
is quite large for a 1.0 BCFD feed gas rate and here is 18 foot in
diameter, which currently is close to the limit of what can be constructed
within a reasonable cost. The bottom section 222 of the column is much
smaller and has a diameter of 9 feet. The bottom section 222 would include
10-15 theoretical trays, equivalent to 18-24 actual trays; the top section
224 includes three theoretical trays, equivalent to four actual trays.
The example process produces an incremental ethane recovery of 12.4%, i.e.
67.4% as compared to 55% in a conventional plant; and incremental propane
and recovery of between 2-3%, i.e. 98.4% as compared to 96.1%. Higher
ethane recoveries would be attainable by increasing the size of the heat
exchangers. This would cool the inlet gas to a greater degree or extent
than if less heat exchange is provided, effectively increasing ethane flow
down the column 220 and reducing the amount of ethane carried over with
residue gas. Increasing recompression of the residue gas could draw down
the temperature and pressure in the column, also leading to increased
recoveries.
It can also be noted that a plant following the Ortloff design in U.S. Pat.
No. 4,278,457 cannot be accomplished in a single column, since
conventional engineering design would require the top section to be at
least a 30 ft height with a 17 ft diameter. This could not practically be
supported on top of a 9 ft diameter bottom section.
This particular configuration would be able to produce a range of ethane
recoveries from 50% up to about 95% by increasing the horsepower of
compressor 125 to reduce the pressure of the columns 220. Above 95%
recovery the pressure in the distillation column would drop to much lower
values and flooding may occur in the top section 224 of the column,
thereby limiting the upper end of the recovery efficiency.
Residue gas leaving the top of the distillation column 220 is sent through
exchanger 206, cooling the gas flow, and is subsequently sent through heat
exchange elements 202 and 200, providing cooling for the inlet gas. This
warms the residue or overhead gas up to 56.degree. F. (at 185b) which is
only slightly colder than the inlet gas temperature. The warm gas is then
compressed at 120 to a pressure of 358 p.s.i.a using the power developed
by the expander 80. The compressed gas is further compressed at 125 up to
a high enough pressure to put it back into the main gas transmission line.
Gas leaving the recompressor 125 is cooled by water or air in heat
exchanger 130 and is further cooled by a heat exchanger 212 with upstream
residue gas. In general, the residue gas is required to leave the plant at
the same temperature at which is entered.
The following Table 3 sets out process parameters that would be obtained in
the plant of FIG. 5a, again based on a theoretical simulation of the
plant.
TABLE 3
Temperature Pressure Molar Flow Vapour
Location .degree. F. p.s.i.a lb moles/hr Fraction
INLET STREAM 68 598.28 116546.84 1
30
STREAM 40a 67.87 596.25 24596.25 1
stream 40b -86.82 590.74 24596.25 0.95
MAIN STREAM 67.87 596.25 91950.59 1
210
stream 35 -93.35 590.74 80970.13 0.93
STREAM 160a -36.0000 593.49 10980.47 1
stream 160b -137.19 578.49 10980.47 0
stream 160c -152.14 329.83 10980.47 0.12
stream 35 -93.35 590.74 80970.13 0.93
stream 62 -91.93 590.74 105566.38 0.94
stream 75b -122.27 331.83 6780.64 0.23
stream 70 -92.28 587.69 98785.74 1
stream 73 -132.76 329.83 98785.74 0.94
STREAM 115 -140.57 327.83 111113.67 1
stream 185a -112.94 322.83 111113.67 1
stream 185b 56.2 320.02 111113.67 1
stream 185c 99.94 313.49 111113.67 1
stream 185d 122.89 358.67 111113.67 1
stream 185e 223.99 626.71 111113.67 1
stream 185f 109.94 621.63 111113.67 1
stream 185g 69.08 613.51 111113.67 1
Reference will now be made to FIG. 5b, which shows a further variant of the
apparatus and method of the present invention. This again is for 1000
MMSCFD plant, and is similar in many respects to the configuration shown
in FIG. 5a. For this reason, and again as before, description of common
components is not repeated, and these common components are given the same
reference numeral.
The difference in FIG. 5b is, in effect, that the two streams 73, 160b of
FIG. 5a are now combined into a single stream (160c), before entering the
top of the fractionation column 220. Thus, in FIG. 5b, the stream 160b
joins with stream 73 in a static mixer 214. The combined stream downstream
from the static mixer 214 is indicated at 160c. The static mixer 214 is a
motionless device, i.e. without any moving parts, and of known
construction. It causes swirling in the fluid flowed downstream, which
increases the turbulence in the piping and consequently increases the mass
transfer rate. If the static mixer has a sufficient length, it is possible
to approach one theoretical stage of contacting, equivalent to one
theoretical stage in the column 220. It is expected that, for situations
where it is not practical or possible to put trays in the top of the
column 220 (e.g. in some retrofit applications where demethanizer columns
are not belled or enlarged at the top at all) this configuration should
provide a significant improvement at even less cost than retrofit with
trays.
It can be noted that, as detailed above, the three tray modification gave
an ethane recovery of 67.4%, whereas, as detailed in Table 4 below, static
mixer 214 gives a recovery of 65.2%. Again, these are theoretical
simulated results only and are predicated on the assumption that a full
equilibrium stage can be achieved in the static mixer 214. It is expected
that actual test results would be close, although likely not quite so
good.
A variant of the proposal of FIG. 5b is shown in FIG. 5c. Here, there is
provided a separator tank or vessel 215, similar to the separator 65, to
separate the gas and liquids from the static mixer 214. Such a separator
would have an outlet for liquid connected to the top of the column 220, as
shown for stream 160c, and an outlet for vapour or gas connected to the
stream 115, and indicated at 160d. Thus, FIG. 5c shows a typical
application of a static mixer modification in which the lighter or vapour
portion of the combined streams 73 and 160b is separated in the separator
215 and then flows around the demethanizer column 220 through the branch
stream 160d; as this vapour flow can be a significnat part of the total
flow, this can significantly reduce the flow rates at the top of the
column 220. Theoretical simulations of the improvement offered by the
static mixer 214 indicated that an 8% incremental ethane recovery could be
achieved by implementing static mixing in the manner shown. Again, this is
under the assumption that one theoretical stage of mixing could be
achieved in the static mixer. With proper design and contact time in the
mixer it should be possible to closely approach the predicted performance
and effect of the static mixer 214.
TABLE 4
Temperature Pressure Molar Flow Vapour
Location .degree. F. p.s.i.a lb moles/hr Fraction
INLET STREAM 68 598.28 116546.84 1
30
STREAM 40a 67.87 596.25 24596.25 1
stream 40b -90.97 590.74 24596.25 0.94
MAIN STREAM 67.87 596.25 91950.59 1
210
stream 35 -92.57 590.74 80970.13 0.93
STREAM 160a -36 593.49 10980.47 1
stream 160b -134.41 578.49 10980.47 0
stream 160c -138.59 330.06 109643.09 0.89
stream 62 -92.2 590.74 105566.38 0.93
stream 75b -122.56 332.06 6903.75 0.23
stream 70 -92.56 587.69 98662.63 1
stream 73 -132.93 330.06 98662.63 1
STREAM 115 -138.74 328.06 111205.24 1
stream 185a -111.5 322.86 111205.24 1
stream 185b 55.97 320.25 111205.24 1
stream 185c 99.82 313.72 111205.24 1
stream 185d 122.67 358.7 111205.24 1
stream 185e 223.53 626.71 111205.24 1
stream 185f 109.94 621.63 111205.24 1
stream 185g 67.98 613.51 111205.24 1
Reference will now be made to FIGS. 6 and 7, which show details of the
trays as shown in the top of the column 220. As shown, the column 220 has
a relatively narrow bottom or body section 222 and a top or upper section
224 of much greater diameter. The bottom section 222, as indicated
includes 10-15 theoretical trays, equivalent to 18-24 actual trays.
In the top section 224 of the column, there are three additional trays 13,
14 and 15, here identified as 226, 227 and 228.
The lowermost of the three trays, tray 226 has a central horizontal portion
230, with upwardly extending lips 232 along opposite sides, which lips 232
provide a weir to maintain a desired fluid level on top of the central
portion 230. In known manner, the central portion would be provided with
an array of bubble caps or valves. Side walls 234 include partially
inclined sections 235, which are connected to a lower floor 236. Again, in
known manner, the floor 236 can be provided with slots, permitting fluid
to flow down through the slots to the tray below.
The tray 227 has a generally circular horizontal portion 240 and is
provided with a slot 241 extending diametrically. Side walls 244 define
the slot and include lips 242, again providing a weir function. The side
walls 244 include lowermost inclined sections 245, to which a floor 246 is
connected. Again, the floor 246 would be provided with slots for downflow
of liquid, while the horizontal portion 240 would be provided with bubble
caps or valves, to permit upward flow of vapour.
The uppermost tray 228 corresponds in many ways to the tray 226, and
includes a central horizontal portion 250, lips 252 and side walls 254.
The side walls 254 again include inclined sections 255, but here these
incline outwardly and away from one another; for the tray 226, the
inclined sections 235 incline inwardly, to follow a frusto-conical
transition section 225 between the top and bottom sections 222, 224.
It will be appreciated that detailed design of the trays, bubble caps,
louvres, mounting arrangements etc., can be largely conventional and
follow known design practice. Such details do not form part of the present
invention.
As noted above, conventional fractionation columns have a disengagement
section, which typically may be 10 to 15 feet in height, with a diameter
of 10 to 18 feet. Here, the top of the column has an upper cap 260. The
distance between the lower edge of the upper cap 260 and the horizontal
portion 250 of the tray 228 is only 2 foot 4 inches. It has surprisingly
been found that this gives adequate disengagement or separation of the
vapour and liquid phases.
There are also other benefits associated with the present invention.
Significantly, the carbon dioxide content of the residue gas stream 115 is
reduced, which aids in alleviating the CO.sub.2 freezing problems commonly
encountered at the upper sections of the column. Instead, a higher
proportion of the carbon dioxide gas in the inlet gas is recovered in the
ethane plus stream leaving the bottom of the column. The reason for the
increased recovery of carbon dioxide is that it has a boiling point quite
close to the boiling point of ethane. Therefore, as there is a higher
recovery of ethane in the present invention, there is also a higher
recovery of carbon dioxide. In trials, the carbon dioxide content of the
residue gas leaving the upper portion of the column was reduced from 0.54
mole % to 0.52 mole %, at the same inlet gas composition.
Additionally, because the present invention requires the addition of the
reflux section 150, there is an increased gas processing capability, since
the inlet gas is initially further divided, and there is less gas flowing
through the inlet heat exchangers.
It will be understood by persons skilled in the art that, although the
present invention has been described in relation to a particular system,
the invention may be implemented in any number of ways, particularly in
the manner in which the various hydrocarbon streams are processed and
delivered to the fractionation column. All such modifications are
contemplated by the present invention.
Additionally, while the present invention has been described largely as a
retrofit to existing facilities, it will be appreciated that the invention
may be utilized in new facilities, and that advantages will be realized in
such use. In particular, an increased recovery of ethane plus may be
substantially achieved over conventional reflux designs, by the
elimination of the high capital cost of a second separator which has
conventionally been required.
It will be appreciated that various changes and modifications may be made
within the spirit of the invention, and all such changes are included
within the scope of the attached claims.
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